UA46176C2 - METHOD OF SEPARATION OF METHANE-CONTAINING GAS FLOW, C ₂ - COMPONENTS, C ₃ - COMPONENTS AND HEAVY HYDROCARBON COMPONENTS - Google Patents

Abstract

A method of ethane, ethylene, propane, propylene, and more substantial carbon components extraction from a hydrocarbon flow is described. In the method described a system for a secondary boiling of the fractioning tower is modified in such a way that one or several liquid distillation tower flows are used, outputting at higher level than it is usually used in the ordinary secondary boiling, providing the partly cooled flow (flows) in the tower to the heater (heaters) delivering, and as a result a more effective cooling of the feeding flows is performed, an by that the more effective of the components needed producing is performed. More over, the liquid flows in the tower outputted at a more high place contain more volatile components, providing better separating from the unwanted components such as carbon dioxide, not lowering the extraction of required components. The hot distillation flow is directed to the lower part of the fractionating tower, at a distance equal to at least one theoretical fractionating stage.

Description

Description of the invention

The present invention relates to a method of separating gas containing hydrocarbons.

Ethylene, ethane, propylene, propane and/or heavier hydrocarbons can be produced from many gases such as natural gas, refinery gas, and synthetic gas streams derived from other hydrocarbons such as coal, crude oil, heavy gasoline, oil shale , tar sands and lignite. As a rule, the majority of natural gas is methane and ethane, that is, methane and ethane together make up at least 5Omol.oo of natural gas. This gas also contains relatively smaller amounts of heavier hydrocarbons such as propane, butanes, pentanes, and the like, as well as hydrogen, nitrogen, carbon dioxide, and other gases.

The present invention generally relates to the extraction of ethylene, ethane, propylene, propane and heavier hydrocarbons from streams of such gases. The structural and group analysis of the gas stream, which is processed according to this invention, in molar percentages is approximately as follows: 92.1295 methane, 3.9695 ethane and other C2-components, 1.0595 propane and other Co-components, 0.1595 isobutane, 0.2195 normal butane, 0.1195 pentanes, the rest - 72 nitrogen and carbon dioxide. Sometimes sulfur-containing gases are also present.

Historically, cyclical fluctuations in the prices of both natural gas and natural gas liquids have reduced the increasing value of ethane, ethylene, propane, propylene and heavier liquid components over time. The fight for processing rights has forced plant operators to maximize the technological capabilities and production efficiency of their existing gas processing plants. The currently known methods of separation of the mentioned substances are based on gas cooling, oil absorption and absorption of cooled oil.

In addition, cryogenic methods have become popular due to the availability of economical equipment that produces energy and simultaneously expands the gas being processed and removes heat from it. Depending on the pressure of the gas source, the enrichment (the content of ethane, ethylene and heavier hydrocarbons) of the gas and the given end products, each of these methods or a combination of them can be applied. with

Today, preference is given to the method of extracting liquid products of natural gas by the method of cryogenic Ge) expansion, because this method provides maximum simplicity, easy start-up, operational flexibility, good productivity, safety and high reliability.

Способи, що мають відношення до цього питання, описуються у патентах США МоМо4175904, 4171964, 4185978, 4251249, 4278457, 4519824, 4617039, 4687499, 4689063, 4690702, 4854955, 4869740, 4889545, о 5275005, 5555748, 5568737, 5771712, 5799507, 5881569, 5890378, reissued US patent Mo33408 co-pending application Mo09/054802, (although the description of this invention in some cases is based on processing conditions that differ from the conditions described in the listed patents and patent applications). --

In a typical cryogenic expansion process, a feed gas stream under pressure (He) is cooled by heat exchange with other streams involved in the process and/or by external cooling sources, such as a propane compression-cooling unit. During the cooling of M gas in one or more separators, liquids may condense and collect in the form of high-pressure liquids containing some of the specified C components. Depending on the gas enrichment and the amount of liquids formed, these high-pressure liquids can be expanded to a lower pressure and fractionated. "

Evaporation, which occurs during the expansion of the mentioned liquids, leads to further cooling of the C 70 stream. Under certain conditions, pre-cooling of high-pressure fluids may be necessary before expansion to further reduce the temperature resulting from expansion. The expanded flow, which is a mixture of liquid and steam, is fractionated in a distillation (methane-distillation) column. In this column, the cooled expansion stream(s) are distilled to separate residual methane, nitrogen and other volatile gases in the form of overhead steam from the specified C»o components, C3 components and heavier hydrocarbon components in the form of a bottom liquid product. shk If the feed gas condenses incompletely (usually it does not condense completely), at least

Gee! part of the vapor remaining after the partial condensation can be passed through the working expansion machine or the engine or the expansion valve to reduce the pressure at which additional liquids are condensed as a result of further cooling of the flow. The pressure after expansion is practically equal to the pressure at which the distillation column is activated. The combined vapor-liquid phases formed as a result of expansion are fed into the column as feed. In recent years, the preferred methods of hydrocarbon separation involve the introduction of such an expanded vapor-liquid flow at the feed point into the middle part of the column, while the upper, absorption section provides additional rectification of the vapor phase

The source of the irrigation stream for the upper rectification section is usually a part of the steam mentioned above, 29 which remains after the partial condensation of the feed gas, but which is removed before the working expansion

GF) Another source for the upper irrigation flow can be a recirculated residual gas flow, which is supplied under pressure. Regardless of the type of source, this vapor stream is typically cooled to significant condensation as a result of heat exchange with other streams involved in the process, such as the cold top of the distillation column. Some or all of the high-pressure liquid resulting from the partial condensation of the feed gas may be combined with said vapor stream prior to cooling. The resulting significantly condensed stream is then expanded by passing through a suitable expansion device, such as an expansion valve, to the pressure at which the methane stripping column is actuated. During expansion, part of the liquid evaporates, which leads to cooling of the general flow. The rapidly expanded stream is then fed to the methane stripping column as overhead feed. As a rule, the steam part of the expanded flow and the steam of the upper run of the methane separation column are combined in the upper, separation section in the rectification column in the form of residual methane product gas. Alternatively, the cooled and expanded stream can be fed to a separator to form the vapor and liquid streams, so that the vapor then joins the top of the distillation column and the liquid enters the column as the overhead feed.

Fig. 1, which depicts a scheme of the technological process, shows the design of a technological installation for the extraction of Co-components from natural gas by a known method according to the US patent Mo4157904. Since this is a large installation, designed for 1.0 billion cubic feet (28,316,900 О0Odm U) of feed gas per day, the methane stripping column (rectification column) had to be designed in two sections with an absorption column 17 and a stripping column 19. In this reproduction of the process, the inlet gas enters into the plant as a 70 31 stream with a temperature of 86"P and a pressure of 613 psi (42.91 kg/cm). If this inlet gas contains a concentration of sulfur compounds that does not allow the product streams to meet the technical conditions, such sulfur compounds are removed by an appropriate by pretreatment of the feed gas (not shown). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. A solid desiccant is usually used for this purpose

The feed flow 31 is cooled in the heat exchanger 10 by heat exchange with the cooled residual gas with a temperature of -99"E (flow 37a), the boiler liquids at the base of the methane column having a temperature of 31"E (flow 42), the liquids of the lower side boiler of the methane distillation column having a temperature of -5"E (stream 41) and liquids of the upper side boiler of the methane expansion column having a temperature of -99"E (stream 40). It should be noted that in all cases the heat exchanger 10 is either a series of individual heat exchangers, or a single multi-pass heat exchanger, or any combination thereof (The decision to use more than one heat exchanger for said cooling will depend on a number of factors including, (but not limited to, volume flow rate of the inlet gas, size of the heat exchanger, flow temperatures, etc.). The cooled stream Z1a enters the separator 11 with a temperature of -82"E and a pressure of 603 psi (42.21 kg/cm), where the vapor (stream 32) is separated from the condensed liquid (stream 35). Ha

The steam from separator 11 (stream 32) is separated into two parts, stream 33 and stream 34. Stream 33, containing about 1895 of the entire steam, is combined with the condensed liquid from separator 11. The combined stream 36 passes i9) through heat exchanger 12, exchanging heat with the steam flow 37 of the upper shoulder of the methane separation column, which leads to cooling and significant condensation of this flow. The highly condensed -139"E stream is then rapidly expanded by passing through a suitable expansion device, such as expansion valve 13, to the operating pressure (approximately 333 psi (23.31 kg/cm2)) of the absorption column 17 of the rectification column. During expansion, part of the flow evaporates, as a result of which the total flow is cooled. In the process shown in Fig. 7, the expanded flow 360, leaving the expansion valve 13, reaches a temperature of -151"E and enters the separation section 17a in the upper part of the absorption column 17. The liquids separated here become the overhead feed for the theoretical stage 1 in the i-o rectification section 1765 (The dashed line indicates another path for the separator liquid (stream 35) according to US Pat. No. 4,278,457, by which at least part of this expanded to a pressure of approximately 333 psi (23.31 kg/cm") by expansion valve 16, cooling stream 35 and forming stream ZbBa, which then goes to the rectification section in the absorption column 17 at the bottom feed " or to the stripping column 19 at the top feed).

The remaining 8295 steam from separator 11 (stream 34) passes to the working expansion machine 14, in which mechanical energy is extracted from this six-sixth part of the high-pressure feed gas. 42.21 kg/cm?) to a pressure of about 333 psi (23.31 kg/cm), with this working expansion cooling the expanded C4a stream to a temperature of about -125°F. Typical commercially available expanders are capable of regenerating 80-8595 of the work theoretically possible with ideal isentropic expansion. The recovered work is often used to drive a centrifugal compressor (such as one designated 15) which could be used to re-compress, for example, the residual gas (stream 37c). The expanded and partially condensed stream 344 is fed to a distillation column in in the place of the lower power supply (in this case - below the theoretical tier 7). -I 20 Liquids (stream 38) from the lower part of the absorption column 17 with a temperature of -127"E are fed by means of a pump 18 into the stripping column 19 at the top feed (stream Zva). The working pressure of the stripping column 19 (343 pounds per sq. inch (24.01kg/cm?)) is slightly higher than the operating pressure of the absorption column 17 and this pressure difference between the two mentioned columns provides the driving force that allows the top steam (stream 39) at a temperature of -125"E to flow from the top of the separation column 19 in the place of the lower feed of the absorption 59 column 17.

GF) The methane stripping section in the absorption column 17 and the stripping column 19 is a conventional type 7 distillation column containing a number of vertically spaced plates, one or more nozzles or a combination of plates and a tower nozzle. The absorption column can consist of two sections, as is often the case in natural gas processing plants. The upper section 17a is a separator in which the partially vaporized overhead feed 60 is separated into its respective vapor and liquid portions and in which the vapor rising from the lower distillation or rectification section 170 is combined with the vapor portion (if any) upper feed for the formation of a distillation stream 37 of cold residual gas, which comes from the upper part of the column. The bottom, rectification section 176 and stripping column 19 contain plates and/or a tower nozzle and provide the necessary contact between the liquids falling down and the vapors rising up. The stripping 65 column 19 also includes reboilers that heat and vaporize portions of the liquids flowing into the bottom of the column to create desorbing vapors that rise to the top of the column.

The liquid product (stream 43) exits the bottom of the column at a temperature of 43°C, based on the fact that the standard technical condition for the bottom liquid product is a molar ratio of methane to ethane of 0.0237 1 and is pressurized to approximately 550 psi .inch (38.5 kg/cm 2 ) (flow 43a) by pump 20. (The pump outlet pressure is generally determined by the final destination of the liquid product. Typically, the liquid product enters storage and the pump outlet pressure is set to prevent any -the evaporation of the flow 43Za as it is heated to the temperature of the surrounding medium).

The residual gas (flow 37) passes countercurrently relative to the inlet feed gas in: a) heat exchanger 70 12, in which it is heated to -99"E (flow 37a), b) heat exchanger 10, where it is heated to 79"E (flow 375) , and c) heat exchanger 21, where it is heated to 110"E (flow 37s) The residual gas is further re-compressed in two stages. In the first stage - with the help of the compressor 15, which is driven by the expansion machine 14, and in the second stage - with the help of compressor 22, driven by an auxiliary power source.

After cooling stream C7e to a temperature of 135" (stream 377) by cooler 23 and to 86"E by 75 with heat exchanger 21, the residual gas product (stream 379) enters the main gas pipeline at a pressure of 631 psi (44.17 kg /cm?), which meets the requirements for the pressure of the gas pipeline (as a rule, this is approximately the pressure at the inlet).

The following table shows aggregated data on volumetric flow rates and energy consumption for the method shown in Fig.1. (fig)

Aggregate data on volumetric flow rates (pound-molecules/hour) can be found below; m7 0000990 12 tvv o za| vovo zvo at 120 here 1ozov2 o zvo zvo zvo 12/10) zvo m -

F in! ' - : « mining (data are based on unrounded values of volumetric flow velocities)

Ethane 84.8995

Propane 96.9096 "

Butanikhk 99.3395 - 70 Power (in horsepower) s Compression of residual gas 44408 2" more pops

The known process, the scheme of which is shown in Fig. 1, gives 84.8995 ethane production when using the appropriate power to compress the residual gas (maximum 45,000 horsepower). However, in the ethane product (a stream containing methane, ethane and carbon dioxide, and is formed, when the bottom liquid product is further fractionated to separate the C-components and lighter components from the C-3-components and heavier components

Ф of hydrocarbon components) the concentration of carbon dioxide is 7.59 moles, which exceeds the set limit set by the owner of the installation, which is b.0 moles maximum. 20 feasible There are many methods of carbon dioxide removal (inlet feed gas treatment, total liquid product treatment, ethane product treatment after fractionation, etc.), but all these methods will increase not only the installation cost (due to the installation costs of the mentioned treatment system), but and operational costs of the installation (due to the consumption of energy and chemicals in the mentioned treatment system)

One way to keep the concentration of carbon dioxide in the ethane product within acceptable limits is to ensure the operation of the methane stripping column in such a way as to drive away carbon dioxide from the bottom liquid

HF) of the product, adding more heat to the column, which is wasted for evaporation, using side reboilers and/or bottom reboilers. and Fig. 2 shows such an alternative working scheme for the process of Fig. 1. The process from Fig. 2 was carried out with the same feed gas composition and conditions as for the described process from Fig. 1. However, when reproducing the process 60 of fig. 2 operating conditions were set to regulate the temperature in the lower part of the stripping column 19 to keep the carbon dioxide content in the ethane product within the technical requirements.

When reproducing this process, as well as when reproducing the process from Fig. 1, the operating conditions were chosen such that the level of ethane production remains as high as possible, while the corresponding power for compressing the residual gas is not exceeded. The feed flow 31 is cooled in the heat exchanger 10 in bo as a result of heat exchange with cold residual gas with a temperature of -96"E (flow 37a), boiler liquids at the base of the methane separation column with a temperature of 50"E (flow 42), liquids of the lower side boiler methane stripping column with a temperature of 38"E (stream 41) and liquids of the upper side boiler of a methane stripping column with a temperature of -32"E (stream 40). The cooled stream 31a enters the separator 11 with a temperature of -72"E and a pressure of 600 pounds per square inch (42kg/cm?), in which the vapor (stream 32) is separated from the condensed liquid (stream 35).

The vapor (stream 32) from separator 11 is separated into two streams, stream 33 and stream 34. Stream 33, containing about 1795% of the entire vapor, is combined with the condensed liquid from separator 11. The combined stream 36 passes through heat exchanger 12, exchanging heat with the steam flow 37 of the upper shoulder of the methane separator 70 of the Column, which leads to cooling and significant condensation of this flow. The significantly condensed Zba flow with a temperature of -132"E is then rapidly expanded by passing through the expansion valve 13. As the flow expands to the working pressure absorption column 17 (326 psi (22.82 kg/cm 2 )), it is cooled to about -152°E (stream 365) Expanded stream 366 enters the column as an overhead feed.

Another portion (remaining 83905) of steam from separator 11 (stream 34) enters an operating expansion machine 14 in which 75% of this high-pressure feed portion extracts mechanical energy. Machine 14 expands the steam primarily isentropically from a pressure of approximately 600 psi ( 42 kg/cm") to the operating pressure of absorption column 17 (326 psi (22.88 kg/cm 2 )), with the operating expansion cooling the expanded stream C4a to a temperature of approximately -118"RH. Expanded and partially condensed stream 344 enters the distillation column as feed at point 11 of the lower feed. 20 Liquids (stream 38) from the lower part of the absorption column 17 with a temperature of -120"E are supplied by means of a pump 18 to the stripping column 19 at the top feed (stream Zva). The working pressure of the stripping column 19 (336 pounds per square inch (23, 52 kg/cm )) is slightly higher than the operating pressure of the absorption column 17, and this pressure difference between the two mentioned columns provides the driving force that allows the vapors of the top runner (stream 39) with a temperature of -118"E from the top of the stripping column 19 drain to the place of the lower feed in the absorption column with 25 17. o

The liquid product (stream 43) exits the bottom of column 19 at a temperature of 56°C. By means of pump 20, the pressure of this stream is brought to about 550 psi (38.5 kg/cm 2 ) (stream 43Za). The residual gas (stream 37 ) passes countercurrently relative to the incoming feed gas in a) heat exchanger 12, where it is heated to -96 E (flow 37a), b) heat exchanger 10, where it is heated to 70"E (flow 375) and c) heat exchanger 21, where ee, 30, it is heated to 101" (flow 37c). The residual gas is then re-compressed in two stages with the help of the rch-compressor 15, which is driven by the expansion machine 14, and with the help of the compressor 22, which is driven by the auxiliary power source. After cooling stream Z7e to 1157E (stream 371) - with the help of cooler 23 and to 86"E with the help of heat exchanger 21, the residual gas product (stream 374) enters the main gas pipeline with a pressure of 631 pounds per square inch (44 17kg/cm ).

The following table shows aggregated data on volume flow rates and energy consumption for the M method shown in Fig. 2.

Table I! : 40 Cumulative data on volumetric flow rates (pound-molecules/hr) are . » g

F

- this 42)

Extraction (data based on unrounded values of volumetric flow rates)

Ethane 68.9495

Propane 96.6195 (F; Butane 99.2595 ke Power (in horsepower)

Compression of residual gas 44641 60

The concentration of carbon dioxide in the ethane product in the method of Fig. 2 is 5.95 mol 905, which corresponds to the technical requirement for such an installation, 6.0 mol 95 maximum. However, it should be noted that the molar ratio of methane to ethane in the bottom product is 0.0008:1 against the permissible ratio of 0.0237:1, which indicates the degree of excessive distillation of light fractions, which was necessary to control the content of carbon dioxide in the liquid products to a given level. Comparison of extraction levels shown in Tables | and 1, shows that the implementation of the method of Fig. 2 in this way in order to reduce the content of carbon dioxide in the ethane product causes a significant reduction in the extraction of liquids. The process of Fig. 2 reduces ethane production from 84.8995 to 68.9495, propane production from 96.9095 to 96.6195, and butane production from 99.3395 to 99.25905.

In the process from Fig. 2, there are two factors that lead to a decrease in the extraction of liquids from the lower part of the separator

Columns 19 compared to the process from Fig.1. First, as the temperature at the bottom of the stripping column 19 rises from 43" in the process of Fig. to 56" in the process of Fig. 2, these temperatures at each location in this column are increased relative to their respective values in the process of Fig. 1 . This reduces the degree of cooling that the liquid streams of the column (streams 40, 41 and 42) can give to the feed gas in the heat exchanger 10. As a result, the cooled feed stream (stream 31 a) entering the separator 11 is warmer (-72"E for the process of Fig. 2 7/0 versus -82"E for the process of Fig. 1), which, in turn, leads to a lower retention of ethane in the absorption column 17, which is reflected by the content of ethane in stream 38 (from 841 lb- molecules/hr for the process of Fig. 2 versus 4734 lb-molecules/hr for the process of Fig. 1). Second, higher temperatures in stripping column 19 force higher temperatures in absorption column 17, resulting in less liquid methane entering stripping column 19 (6,842 lb-mol/hr in stream 38 for the process of Fig. 2 vs. 11021 lb-mol/h /5 for the process of Fig. 1) When this liquid methane is further vaporized by the side reboilers and the main reboiler attached to the stripping column 19, the methane vapor helps drive off carbon dioxide from the flowing liquids down the column. When reducing the methane involved in the process from fig. 2 for the distillation of carbon dioxide, more ethane in these liquids must evaporate to serve as a desorbing gas.

Since the relative volatilities of carbon dioxide and ethane are very similar, ethane vapor is a much less effective 2o desorbing agent than methane vapor, which reduces the efficiency of hydrocarbon recovery in the column.

The task of this method is to carry out a separation, as a result of which the residual gas is obtained, which is formed at the end of the process and which contains practically all the methane in the feed gas and contains almost none of the CO components and heavier hydrocarbon components, and the lower fraction that comes from metal separation column, contains almost all Co-components and heavier hydrocarbon components and almost does not contain methane or more volatile components, while meeting the technical conditions regarding the maximum permissible content of carbon dioxide.

In the first variant, the means of separation of the gas flow containing methane, Co-components, C3-components and heavier hydrocarbon components into a volatile fraction of the residual gas, which contains most of the methane, and a relatively less volatile fraction, which contains most of the Co-components , C3-components and heavier Hezo Hydrocarbon components, according to which: the gas stream is processed by subjecting it to one or more stages of heat exchange and at least one - stage of separation, to form at least the first feed stream, cooled under pressure for its almost complete condensation , and at least a second pressure-cooled feed stream, the substantially condensed first feed stream is expanded to a lower pressure, causing it to flow even more)

Zv is cooled and fed to the rectification column at the upper feed point, "E cooled second feed stream is expanded to a lower pressure and fed to the rectification column at the feed point in the middle part of the column, cooled expanded first feed stream and expanded second feed stream are fractionated at more low pressure, as a result of which the components of the relatively less volatile fraction are driven away, "according to the invention, the problem is solved by the fact that: in c, the liquid distillation stream is removed from the rectification column and heated, the heated distillation stream is returned to the place in the lower part of the rectification column, separated by from the place of removal by at least one theoretical tier, withstand the values and temperatures of the feed streams entering the rectification column, sufficient to maintain such a temperature of the upper run of the distillation column, at which most of their components in the relatively less volatile fraction are driven away.

In the second version of the method, according to which:

Me. the volatile fraction of the residual gas is re-compressed and a part is removed to form the compressed first feed stream, the compressed first feed stream is cooled under pressure for its almost complete condensation, the practically condensed first feed stream is expanded to a lower pressure, resulting in

Ф it is further cooled and fed to the distillation column at the top feed, the gas stream is subjected to one or more stages of heat exchange to form a second feed stream cooled under pressure, 5B the cooled second feed stream is expanded to a lower pressure and fed to the distillation column in the feeding place in the middle part of the column,

F) the cooled expanded first feed stream and the expanded second feed stream are fractionated at a lower pressure, as a result of which the components of the relatively less volatile fraction are driven off, according to the invention, the problem is solved by the fact that: the liquid distillation stream is removed from the rectification column and heated, the heated distillation stream is returned to the lower part of the distillation column, separated from the place of removal by at least one theoretical layer, the values and temperatures of the feed streams entering the distillation column are maintained, sufficient to maintain such a temperature of the upper run of the distillation column, at which most of the 65 components in relatively less volatile fractions are driven off.

The present invention makes it possible to use a new installation or modification of an existing technological installation for the implementation of a separation method that solves the problem at significantly lower capital costs by reducing or eliminating the need for a product treatment system to remove carbon dioxide. On the other hand, the present invention, regardless of whether it is applied in a new installation or as a modification of an existing technological installation, can be used to extract more

of Co-components and heavier hydrocarbon components in the lower liquid product at a given concentration of carbon dioxide in the feed gas than when using Other technological systems.

According to this invention, it turned out to be possible to extract more than 8496 tons of CO while maintaining the content of carbon dioxide in the lower liquid product within the limits of technical requirements and ensuring almost complete removal of methane into the residual gas stream. The present invention, although applicable at lower pressures and warmer temperatures, has particular advantage in the treatment of feed gases at pressures in the range of 600-1000 psi (42.18-70.Zkg/cm?) or higher under conditions , which require temperatures of the upper shoulder of the column -120"E or colder.

The present invention utilizes a modified bottom-of-column reboiler system that can be used with any type of natural gas liquid recovery system. When using a conventional bottom reboiler or side reboiler in a distillation column, the entire liquid stream flowing down the column is diverted from the distillation column and passed through a heat exchanger, then returned to the column at almost the same point. In the mentioned modified boiler system at the base of the column, part of the column liquid that flows down is diverted through a place located higher in the column, that is, separated from the place of return by at least one theoretical tier. Even though the volumetric flow rate of the liquid may be lower, the liquid is generally much colder and this may have the advantage of increasing recovery or reducing the size of the heat exchanger.

It was found that when applying this invention in known methods of extracting liquid components of natural gas, the extraction of C.o-components and heavier components increases by 1-295 Ha. However, such

The increase in extraction becomes significantly higher when it is necessary to reduce the content of carbon dioxide in the extracted liquid product of natural gas. During the extraction of ethane in a typical installation for extraction of liquid i9) components of natural gas, at least a certain amount of carbon dioxide contained in the feed gas is also extracted, because carbon dioxide in terms of relative volatility, it occupies a position somewhere between methane and ethane. Therefore, with the increase in the extraction of ethane, the extraction of carbon dioxide in the liquid product of natural gas also increases. It was found that when using the proposed modified system of the boiler at the base of the column, it is possible to significantly increase the production of ethane in the liquid product of natural gas compared to the methods where - the systems of the conventional boiler at the base of the column or the side boiler are used, when the column is "re- boil to meet the requirement for the given content of carbon dioxide in the liquid product of natural gas. i-th

For a better understanding of this invention, reference is made to examples and drawings in which: "I

Fig. 3 is a diagram of the technological process showing how the installations shown in Figs. 1 and 2 can be adapted to process natural gas according to the present invention.

Fig. 4 is a diagram of the technological process showing another variant of the adaptation of the installations shown on "

1 and 2, for processing natural gas according to the present invention.

Fig. 5 is a scheme of the technological process, showing how another known method can be applied in the installation - c for the treatment of natural gas according to the present invention. Fig. 6 is a diagram showing the modified boiler system proposed by the present invention at the base of the "" column for the technological installation, and this system includes a thermosiphon system.

Fig. 7 is a diagram showing the modified boiler system proposed by this invention at the base

Columns for a technological installation, and this system includes a pumping system. "d" Fig. 8 is a diagram showing the modified boiler system proposed by the present invention at the base of the column for a technological installation, and this system includes a pumping system. b Fig. 9 is a diagram showing the modified boiler system proposed by this invention at the base of the column for a technological installation, and this system includes a system of separate columns.

Tables summarizing volumetric velocities and flows calculated for typical process conditions are added to the further explanation of the figures described above. In the given tables, the values of volume velocities

Flow rates (in lb-molecules per hour) are rounded to the nearest whole number for convenience. The total flow rates shown in the tables include all non-hydrocarbon components and are therefore generally greater than the sum of the volumetric flow rates for the hydrocarbon components. The temperatures shown are approximate values, rounded to the nearest degree. It should also be noted that the calculations of technological processes performed to compare the processes shown in the figures are based on the assumption that there is no heat input from the surrounding environment (or into it) into the process (or from it). Such an assumption, which is common for specialists, is very reasonable due to the quality of insulating materials available on the market.

Example 1 in Fig. C shows a diagram of the technological process according to the present invention. The composition of the feed gas and the conditions considered when describing the process from Fig. 3 are the same as for the process from Fig. 1. In this regard, the process of Fig. 3 can be compared with the process of Fig. 1 to show the advantages of this invention.

In reproducing the process of Fig. 3, the inlet gas enters at a temperature of 86 °C and a pressure of 613 psi (42.91 kg/cm 2 ) as stream 31. The feed stream 31 is cooled in the heat exchanger 10 as a result of heat exchange with the cold residual gas 65 , which has a temperature of -99"E (stream 37a), the liquids of the boiler at the base of the methane stripping column (stream 42), the liquids of the side boiler of the methane stripping column, which have a temperature of -4"E (stream 41) and the liquids from the lower part part of the absorption column having a temperature of -287E (stream 45.) The cooled stream Z1a enters the separator 11 at a temperature of -84"E and a pressure of 603 psi (42.21 kg/cm2), in which the vapor (stream 32) is separated from condensed liquid (stream 35).

The steam (stream 32) from the separator 11 is divided into two gaseous streams, stream 33 | stream 34 Stream 33, containing about 1995% of all steam, combines with the condensed liquid (stream 35) to form stream 36

The combined stream 36 passes through the heat exchanger 12, exchanging heat with the cold residual gas (stream 37), where it is cooled to -138"E. The resulting significantly condensed stream Zba is further rapidly expanded by passing through a suitable expansion device, for example, an expansion valve 13, 70. to the operating pressure (approximately 332 psi (23.24 kg/cm?)) of the absorption column 17. During expansion, a portion of this stream is evaporated, resulting in cooling of the total stream. In the process shown in Fig.3, the expanded stream 365, leaving the expansion valve 13, reaches a temperature of -151 EC and enters the absorption column 17 as the overhead feed. flow 37, 75 which is removed from the upper zone of the column.

Returning to gaseous stream 34, the remaining 81905 of steam from separator 11 enters an operating expansion machine 14 in which mechanical energy is extracted from this high pressure feed portion. Machine 14 expands the steam primarily isentropically from a pressure of approximately 603 psig (42.21 kg/cm?) to a pressure of about 332 psig (23.24 kg/cm?), with this working expansion cooling the expanded stream 3Z4a to a temperature of about -127"E The expanded I! partially condensed stream Z4a then enters the absorption column 17 as feeding at the bottom feeding point of the column.

The dashed line alternatively shows that the condensed liquid (stream 35) from the separator 11 can be rapidly expanded by passing through a suitable expansion device, for example, an expansion valve 16, to the operating pressure of the absorption column 17, cooling the stream 35 and forming a stream Z5Ba Expanded stream ZBa, EM which comes out of the expansion valve 16, can then be fed to the absorption column 17 at the place of its lower d) feed or to the stripping column 19 at the place of its upper feed.

Liquids (stream 38) from the lower part of the absorption column 17 with a temperature of -128"E enter the pump 18, with the help of which their pressure is increased (stream Zva), and are divided into two parts. One part (stream 44) containing approximately 5595 of the entire liquid is fed to the stripping column 19 at the top feed point. with two columns creates a driving force so that the vapors (stream 39) of the upper steam with a temperature of -123"E from the upper part of the stripping column 19 flow to the place of the lower feed of the absorption -- column 17. Ge)

The other part (flow 45), containing the remaining 45905 liquid flow Zva with increased pressure, is sent to the heat exchanger 10, in which it supplies part of the feed gas cooling as it is heated to -20"E and M partially evaporates. The heated flow 45a then enters the of the stripping column 19 at the feed point located in the middle of the column, separated from the top feed point where the stream 44 enters the column by at least one theoretical stage. In this case, the partially evaporated stream flows to the same point on the column that was used for the return of the upper side reboiler (theoretical tier 8 in C 70 stripping column 19) in the process of Fig. 1, which is seven theoretical tiers below the point of removal of the liquid c flow in the fractionation system (the upper feeding place, where flow 44 enters the stripping column 19). z» The liquid product (stream 43) leaves the lower part of the column 19 with a temperature of 42"E. With the help of the pump 20, the pressure of this stream is increased by approximately up to 550 psi (38.5kg/cm?) (4Za flow). The residual gas (flow 37) passes countercurrently relative to the inlet feed gas in a) heat exchanger 12, in which it

В о В В : : о и и 1" is heated to -99"E (stream 37a), b) heat exchanger 10, in which it is heated to 80" (stream 375), and c) heat exchanger 21, in which it is heated to 105"E (flow 37s). The residual gas is then re-compressed in two

Ge) stages, with the help of the compressor 15, which is driven by the expansion machine 14, and with the help of the compressor 22, which is driven by the auxiliary power source. After cooling stream Z7e to 1157 (stream 37) by cooler 23 and to 86"E by heat exchanger 21, the residual gas product (stream 379) enters - 7o main gas line with a pressure of 631 psi (44.17 kg/cm), from where it leaves for implementation.

Ф In the following table, the total data on volumetric flow rates and energy consumption for the method shown in Fig. 3 are given.

Table of PI »

Total data on volumetric flow rates (pound-molecules/hour) schi yuyu 59 bo 45) 53582185 611 277 251) 8698 y - y

Extraction (data based on unrounded values of volumetric flow rates)

Ethane 86.1295

Propane 97.1095

Butanyh 99.4195 70 Power (m horsepower)

Compression of residual gas 44413

A comparison of Tables I and II shows that compared to the known method, this invention increases the extraction of ethane from 84.8995 to 86.1295, the extraction of propane - from 96.9095 to 97.1095, the extraction of butanes - from 99.3395 to 99.4195. A comparison of Tables I and I also shows an increase in performance for equivalent power requirements (use).

When using a modified boiler system at the base of the column, the column liquid flowing into the heat exchanger 710 (stream 45) is colder than the corresponding stream 40 in the process of Fig.1. This enhances the cooling of the inlet gas, so with such a system you can get not only a much higher performance from the liquids, but also use the liquids at a colder temperature level than would be possible when using a conventional boiler system at the base of the column. As a result, in the process of Fig. 3, the extraction of soybean components and heavier hydrocarbon components increases when using almost the same power for compressing the residual gas as in the known process of Fig. 1.

Example 2 p

In those cases where the content of carbon dioxide in the liquid product is a controversial issue (due to more stringent technical requirements for the product, which are set by the customer, such as, for example, in the known method from Fig. 2, described earlier), this invention offers significant advantages in extraction and efficiency compared to the mentioned known method shown in Fig.2. The operating conditions for the process of Fig. 3 can be changed in order to reduce the content of carbon dioxide in the proposed liquid product, as shown in Fig. 4. The composition of the feed gas and the conditions considered when describing the process shown in Fig. 4 are the same as for the processes of Fig. 1 and 2. In this regard, the process of Fig. 4 can be compared with the processes of Fig.1 and 2 to demonstrate the advantages of this invention. -

When reproducing the process from Fig. 4, we used practically the same system of cooling and separation of the inlet gas as in the process from Fig. 3. The main difference was that the plant regulators adjusted to increase the proportion of liquids from the lower part of the absorption column 17 (stream 45), which are heated in the M heat exchanger 10 and flow to the stripping column 19 at the feed point located in the middle of the column.

The regulators of the installation were also adjusted to a slight increase in the temperature in the lower part of the stripping column 19 (from 42"E in the process of Fig. i.e. 0.0237:1. | the increased amount of heated stream 45a entering - 70 the distillation column 19 and the higher temperature of the lower liquids increase the distillation inside the column, due to which the temperatures in the process of Fig. 4 become warmer compared to the process of Fig. 3 and in the absorption column 17, c" and in the stripping column 19, and as a result, the carbon dioxide content of the liquid product, stream 43, leaving the stripping column 19 is reduced. The warmer column temperatures also lead to a slight reduction in cooling, which is transferred from the streams involved in the process and which are used as feed streams for the column.In particular, this requires a slight reduction in the fraction of the separator feed gas (stream 32) that is directed to heat exchanger 12 with stream 33, thus reducing the amount of stream 360 entering the

Ge") absorption column 17 in the place of upper feeding. with In the following table, the total data on volumetric flow rates and energy costs for the process shown in Fig. 4 are given. - 50

IM table;

Cumulative data on volumetric flow rates (lb-mol/h) from о ю во b5

Extraction (data based on unrounded values of volumetric flow rates)

Ethane 84.6195

Propane 96.9695

Butany 99.3995

Power (in horsepower)

Compression of residual gas 44573

The concentration of carbon dioxide in the ethane product of the process from Fig. 4 is 5.80 moles, which is much lower than the requirement set by the customer. A comparison of the recovery levels shown in Tables I and II shows that the present invention allows achieving the required carbon dioxide content while maintaining almost the same liquid recovery efficiency as in the process of Fig. 1. While ethane production decreased slightly from 84.8995 to 84.6190, both propane and butane production increased slightly, from 96.9095 to 96.9695 | from 99.3395 to 99.39905, respectively.

A comparison of Tables 111M also shows that the storage volumes of the mentioned products reached almost t5 with the same needs in (use of) capacity.

A comparison of the extraction levels shown in Tables II and IM shows that the present invention allows for significantly higher liquid extraction efficiency than the method shown in Fig. 2 when it is regulated to limit the carbon dioxide content of the liquid product. Compared to the process of Fig. 2, the process of Fig. 4 increases the extraction of ethane from 68.9495 to 84.6195, that is, by almost 15.795. Also, the extraction of propane and butane increases slightly, from 96.6195 to 96.9695 and from 99.2595 to 99.3995, respectively. A comparison of Tables II and IM also shows that the increase in the volumes of output of the mentioned products was not simply the result of an increase in the demand for (utilization of) capacity. On the contrary, when applying this invention, as in Example 2, not only the extraction of ethane, propane and butanes increases compared to the mentioned known process, but also the extraction of liquid increases by 2395 (based on the amount of ethane extracted per unit of consumed power). high school

As with the process of Fig. 3, a significant advantage achieved by the variant implementation of the process shown in Fig. 4 is that the modified reboiler system at the base of the column utilizes the colder column fluids to cool the inlet feed streams . This increases the cooling that is supplied to the intake gas and it makes it possible in this case to get not only a much higher performance from the liquid, but also at a colder temperature. At the same time, methane is injected into the lower part of the distillation column 19 more than it would be there in the opposite case during re-boiling - the column to reach the specified carbon dioxide content. (It should be noted that stream 45 in the process of Fig. 4 moves 5721 lb-molecules of methane per hour and is introduced at theoretical stage 8 of stripping column 19, while stream 40 in the process of Fig. 2 moves only 1886 lb-molecules of methane per hour and it is introduced into the upper (Se) part of the separation column 19). Additional methane, which according to the present invention is added in the process of Fig. 4, "helps drive away carbon dioxide from liquids flowing down in the drive column. The amount of carbon dioxide in the liquid natural gas product can be adjusted by adjusting the amount of liquid diverted to feed a modified reboiler system at the base of the column instead of feeding the top of the stripping column. « du Fig. 5 shows a diagram of the technological process, showing how the method and installation described in the patent -

US Mo. 5568737, can be converted into a natural gas processing plant according to the present invention. Figures b, 7, 8 and 9 show diagrams showing some other ways of implementing a modified boiler system. Fig. b shows the use of a typical thermosyphon, in which the partial flow of liquid coming from the distillation column 50 to the boiler 57 at the base of the column can be regulated by means of a valve 58 (in the pipeline 61 for liquid removal). The liquid part, not removed from the column, simply flows 395 from the chimney-like plate 51 to the distributor 52 for the tower nozzle (or plates) 53, located below. The heated stream in the pipeline 61 a From the reboiler 57 at the base of the column is returned to (o) the distillation column 50 in a place in the lower part, in which there is a suitable power distribution mechanism - for example, a chimney-like plate 54 and a distributor 55, for mixing the heated flow with column liquids flowing down from the tower nozzle 53 and supplying this mixture to the tower nozzle (or plates) 56. In Fig. 7 and 8 -| 50 shows typical injection means in which all flowing liquid is diverted to the pipeline 61 for discharge

Ф liquid I with the help of pump 60 increases the pressure. The flow of liquid with increased pressure in the pipeline 61a is then divided using the appropriate distribution valves 58 and 59 to direct a given amount of liquid into the pipeline 62 leading to the boiler 57 at the base of the column. The heated flow in the pipeline 62a from the boiler 57 is returned to the rectification column 50 in its lower part, as was described earlier for the variant 59 from fig.b. In the variant from Fig. 7, the liquid that does not enter the boiler (in the pipeline 63),

HF) is returned to the chimney-like plate 51, from which this liquid was initially diverted, as a result of which it

HF can overflow from the chimney-like plate 51 to the distributor 52 and the tower nozzle (or the plate 53 is located below. In the variant of Fig. 8, the liquid that does not enter the boiler (in the pipeline 563) is returned to the area under the chimney-like plate 51 , from which this liquid was initially diverted, directly to the distributor 52, which supplies this liquid to the lower tower nozzle (or plates) 53. Fig. 9 shows how the aforementioned injection system described for Fig. 8 can be applied in the case of separate columns, for example with the upper column 65 and the lower column 50, while the injection system is the same as used in the processes of Fig. 3 and 4.

Those skilled in the art will appreciate that the present invention achieves some of its advantages by supplying 65 a cooler stream to the side reboiler(s) and/or the column bottom reboiler(s), providing additional cooling to the feed or feed columns This additional cooling reduces the amount of power consumed for a given product yield level, or increases product yield levels for a given energy input, or a combination of both. In addition, those skilled in the art will appreciate that the present invention offers the benefit of introducing greater amounts of methane into the lower portion of the methane stripping column, which aids in the distillation of carbon dioxide from the downstream liquids. When the amount of methane required for the distillation of said liquids is increased, the amount of ethane required for distillation is correspondingly reduced, as a result of which it becomes possible to retain more ethane in the lower liquid product. Therefore, the present invention is generally suitable for any process that depends on the cooling of any number of feed streams | feeding the resulting 7/0 feed stream(s) into the column for distillation.

According to the present invention, the cooling of the feed streams of the methane extraction column can be carried out in many ways. In the processes of Fig. 3 and 4, the feed stream 36 is cooled and largely condensed due to the steam stream 37 of the upper collar of the methane stripping column, while the liquids of the methane stripping column (streams 45, 41 and 42) are used only for cooling the gas stream. In the process of Fig. 5, the high-pressure residual gas feed stream 7/5 48 is also cooled and significantly condensed by portions of the vapor stream (streams 46 and 37) of the upper run of the distillation column, while the methane stripping column liquids (streams 40 and 42) are used only for cooling the gas stream However, methane stripping column fluids could be used to provide partial or complete cooling and substantial condensation of stream 36 in Figs. 3-5 and/or stream 48 in Fig. 5 in addition to or instead of cooling the gas stream. In addition to 2o in addition, any stream with a temperature colder than the temperature of the feed stream being cooled can be used. For example, the side steam from a methane stripping column can be diverted and used for cooling. Other potential sources of cooling include (but are not limited to) separator liquids , formed as a result of rapid discharge of high pressure, and mechanical cooling systems Selection of sources and cooling will depend on many factors including (but not limited to) inlet gas composition and conditions, plant size, heat exchanger size, potential cooling source temperature, etc. Specialists in this field of technology also understand that in order to achieve i) a given feed gas temperature(s), it is also possible to use any combination of the above-mentioned cooling sources or cooling methods.

According to the present invention, external cooling can be used in addition to the cooling that is transferred to the inlet gas from other streams participating in the process, in particular in the case when the inlet gas is richer than the inlet gas used in Examples 1 and 2. Use and distribution of methane stripping - column fluids for the heat exchange occurring during the process, the specific arrangement of heat exchangers for "- inlet gas cooling should be determined for each specific application, as should the selection of the streams involved in the process for specific heat exchange functions. ise)

It is not necessary to combine all of the high-pressure liquid (stream 35) in Fig. 3-5 with a portion of the vapor from separator E (stream 33) flowing into the heat exchanger 12. On the other hand, this liquid stream (or part thereof) can be expanded, passing through a suitable expansion device, for example, an expansion valve 16, and fed to the feed point located below the middle of the distillation column (absorption column 17 or distillation column 19 in Fig. 3 and 4, rectification column 17 in Fig. 5). This liquid flow can also be «

Use for cooling the inlet gas or for another heat exchange function before the expansion stage or -ptv) with after it, but before entering the methane stripping column.

It should also be noted that the relative amount of feed contained in each branch of the column feed streams will depend on several factors, including gas pressure, feed gas composition, the amount of heat that can be economically recovered from the feed, and the amount of permissible power . Magnification

Feed to the top of the column may increase recovery, but will reduce the energy that can be extracted from the expander, resulting in increased energy requirements for recompression. Increasing power at the bottom of the column reduces energy costs, but can

Me, to reduce the level of extraction of the product. Preference is given to the places of power supply in the middle part of the column - depicted in Fig. 3 and 4, for the described operating conditions of the processes. However, the relative places of feeding 5o In the middle part of the column can be changed depending on the composition of the inlet gas or other factors, for example - on the given levels of extraction and the amount of liquid formed during the cooling of the inlet gas. Moreover,

It is possible to combine two or more feed streams, or their parts, depending on the relative temperatures and quantities of the individual streams, and this combined stream is then fed to the feed point in the middle part of the column. Figures 3 and 4 show the preferred options for the specified compositions and pressure regime. Although single thread expansion is shown in specific expansion devices, other expansion means may be used where appropriate. For example, conditions can guarantee working

F) expansion of a significantly condensed part of the feed stream (Zba in Fig. 3-5) or a significantly condensed ka stream that is returned for recirculation (485 in Fig. 5).

Figures 3 and 4 show a rectification column that, due to the size of the installation, consists of two sections (17 and 19). The decision to design a distillation column in the form of a single tank (for example, 17 in Fig. 5) or in the form of several tanks will depend on a number of factors, for example, on the size of the installation, the distance to the production equipment, etc.

Although preferred embodiments of the present invention have been described, those skilled in the art will appreciate that modifications are possible, for example to adapt the present invention to different conditions, types of power supply and other requirements, but within the scope of the present invention, the essence of which is defined the formula of the invention.

Claims (13)

The formula of the invention 1. A method of separating a gas stream containing methane, Co-components, C3-components and heavier 2 hydrocarbon components into a volatile fraction of the residual gas containing most of the methane and a relatively less volatile fraction containing most of the Co-components , C3 components and heavier hydrocarbon components, whereby the gas stream 31 is treated by subjecting it to one or more heat exchange steps and at least one separation step to form at least a first feed stream cooled under pressure to substantially 70 completely condense it, and at least of the second feed stream, cooled under pressure, the substantially condensed first feed stream is expanded to a lower pressure, as a result of which it is further cooled, and fed to the rectification column at the top feed, the cooled second feed stream is expanded to a lower pressure and fed to the rectification column the column in the feeding place in the middle part of the column; 12 the cooled expanded first feed stream and the expanded second feed stream are fractionated at a lower pressure, as a result of which the components of the relatively less volatile fraction are driven off, which is distinguished by the fact that the liquid distillation stream is removed from the distillation column and heated; the heated distillation stream is returned in a place in the lower part of the distillation column, separated from the place of withdrawal by at least one theoretical tier, the values and temperatures of the feed streams entering the distillation column are maintained, sufficient to maintain such a temperature of the upper run of the distillation column, at which most of the components in the relatively less volatile fraction is driven away. 2. The method of separation of the gas flow containing methane, Co-components, C3-components and heavier hydrocarbon components into the volatile fraction of the residual gas, which contains most of the methane, and relatively (3) the less volatile fraction, which contains most of the Co -components, C3-components and heavier hydrocarbon components, according to which the volatile fraction of the residual gas is re-compressed and a part is removed to form a compressed first feed stream, and the compressed first feed stream is cooled under pressure for its almost complete condensation, - practically condensed first feed stream the stream is expanded to a lower pressure, thereby cooling it further, and fed to a rectification column at the top feed, -- the gas stream is subjected to one or more heat exchange stages to form a second pressure-cooled feed stream (He) the second feed stream is expanded to a lower pressure and fed into the rectification circle well M at the feed point in the middle part of the column, the cooled expanded first feed stream and the expanded second feed stream are fractionated at a lower pressure, as a result of which the components of the relatively less volatile fraction are driven off, which differs in that the C 50 liquid distillation stream is withdrawn from the rectification column and heated, the heated distillation flow is returned to the lower part of the distillation column, separated from the place of removal by at least one theoretical layer, the values and temperatures of the feed streams entering the distillation column are maintained, sufficient to maintain such a temperature of the upper run of the distillation column, at which most of the components in the relatively less volatile fraction are driven off. t- Z. The method according to claim 1 or 2, which differs in that the liquid distillation stream after its removal from the He») of the rectification column is fed using a pump. 4. The method according to claim 3, which is characterized by the fact that the diluted liquid distillation stream is divided into at least - the first and second parts, the first part is heated and the heated first part is returned to a place in the lower part - and 20 of the rectification column, separated from the point of withdrawal by at least one theoretical level. 5. The method according to claim 1 or 2, which is characterized by the fact that the liquid distillation stream is sent to heat exchange with at least part of the gas stream or feed streams to transfer cooling to them and, thus, to heat the diluted liquid distillation stream. b. The method according to item 3, which is characterized by the fact that the diluted liquid distillation stream is directed to 22 heat exchange with at least part of the gas stream or feed streams to transfer cooling to them and, in this way, to heat the diluted liquid distillation stream. 7. The method according to claim 4, which is characterized by the fact that the first part is directed to heat exchange with at least a part of the gas stream or feed streams to transfer cooling to them and, thus, to the first part. 8. The method according to claim 1 or 2, which is characterized by maintaining the values and temperatures of the heated 60 distillation stream and the heating of the rectification column sufficient to maintain such a temperature in the lower part of the rectification column at which the amount of carbon dioxide contained in the relatively less volatile fraction , decreases. 9. The method according to claim 3, which is characterized by the fact that the values and temperatures of the heated distillation stream and the heating of the rectification column are sufficient to maintain such a temperature in the lower part of the rectification column, at which the amount of carbon dioxide contained in the relatively less volatile fraction, is decreasing 10. The method according to claim 4, which is characterized by the fact that the values and temperatures of the heated first part and the heating of the distillation column are sufficient to maintain such a temperature in the lower part of the distillation column at which the amount of carbon dioxide contained in the relatively less volatile fraction decreases. 11. The method according to claim 5, which differs in that the values and temperatures of the heated distillation stream and the heating of the rectification column are sufficient to maintain such a temperature in the lower part of the rectification column, at which the amount of carbon dioxide contained in the relatively less volatile fraction decreases. . 12. The method according to claim 6, which is characterized by the fact that the values and temperatures of the heated distillation stream and the heating of the distillation column are sufficient to maintain such a temperature in the lower part of the distillation column at which the amount of carbon dioxide contained in the relatively less volatile fraction decreases. 13. The method according to claim 7, which is characterized by the fact that the values and temperatures of the heated first part and the heating of the distillation column are sufficient to maintain such a temperature in the lower part of the distillation column at which the amount of carbon dioxide contained in the relatively less volatile fraction decreases. with 6) (Se) in "- (Se) " - and? sh» (o) - -i 4) ime) 60 b5